Recuperative combustion system

ABSTRACT

The methods and systems described herein relate to a recuperative combustion system that recuperates energy from fuel combustion that would otherwise be lost. The recuperative combustion system minimizes or eliminates the need for an air separator unit through the use of a clean water splitter section, consisting of a thermochemical cycle or high-temperature electrolysis. Water is split into its component hydrogen and oxygen, primarily with process heat from the combustion process. The oxygen produced by the water splitter provides oxygen necessary for oxy-fuel combustion, thereby reducing or eliminating the need for the power intensive air separator unit and/or external oxygen source, significantly increasing the efficiency of the oxy-fuel combustion cycle. Hydrogen produced by the water splitter may be used for a variety of industrial uses, or combined with carbon dioxide (captured from the flue gases produced by said combustion process) to produce methanol. Methanol can further be refined in a methanol to gasoline reactor to produce dimethyl ether, olefins or high grade gasoline. Described herein are methods and systems that 1) increase oxy-fuel combustion efficiency, 2) produce hydrogen for a suite of industrial/energy uses, and 3) capture carbon dioxide and convert it to high value hydrocarbons.

RELATED APPLICATIONS

This application claims priority to U.S. Provisional Application No.61/274,745, filed Aug. 20, 2009, and U.S. Provisional Application No.61/345,541, filed May 17, 2010. The entire contents of these patentapplications are hereby incorporated herein by reference.

FIELD OF THE INVENTION

The methods and systems described herein relate to a recuperativecombustion system that recuperates energy from fuel combustion thatwould otherwise be lost. The recuperative combustion system minimizes oreliminates the need for an air separator unit through the use of a cleanwater splitter section, consisting of a thermochemical cycle orhigh-temperature electrolysis. Water is split into its componenthydrogen and oxygen, primarily with process heat from the combustionprocess. The oxygen produced by the water splitter provides oxygennecessary for oxy-fuel combustion, thereby reducing or eliminating theneed for the power intensive air separator unit and/or external oxygensource, significantly increasing the efficiency of the oxy-fuelcombustion cycle. Hydrogen produced by the water splitter may be usedfor a variety of industrial uses, or combined with carbon dioxide(captured from the flue gases produced by said combustion process) toproduce methanol. Methanol can further be refined in a methanol togasoline reactor to produce dimethyl ether, olefins or high gradegasoline. Described herein are methods and systems that 1) increaseoxy-fuel combustion efficiency, 2) produce hydrogen for a suite ofindustrial/energy uses, and 3) capture carbon dioxide and convert it tohigh value hydrocarbons.

BACKGROUND

Energy supply and concerns over unmitigated greenhouse gas emissions aretwo critical issues of the 21^(st) century. The use of fossil fuels(coal, oil and natural gas) shall continue to play a central role inelectricity production for decades to come, and is projected tosignificantly increase before they are phased out. Hydrogen, a cleanenergy carrier, is anticipated to play a significant role in energyproduction in the future.

While carbon capture and sequestration results in a reduction inatmospheric inputs of carbon from industrial sources, it will requirelarge-scale construction of a pipeline distribution and storage system,which will be extremely costly to build. Additionally, the effectivenessof long-term subterranean CO₂ storage on a scale of hundreds tothousands of years (required for carbon dioxide mineralization) ispresently unproven. Rather than simply disposing of purified, capturedand pressurized CO₂ from oxy-fuel combustion and purification systems atgreat expense, consideration for use of carbon dioxide as an industrialfeedstock for developing reconstituted high-value carbon-based compounds(e.g., hydrocarbons and oxygenated hydrocarbons) may be seen as aneconomically and environmentally attractive alternative. Reconstitutingwaste carbon dioxide into useful materials turns a significant liabilityinto an asset which may 1) reduce carbon emissions and 2) yield fuels,fuel precursors and other beneficial industrial compounds and products.

SUMMARY OF THE INVENTION

In one aspect, provided herein is a method for oxy-fuel combustion,comprising:

providing a system comprising a combustion chamber arranged and disposedto receive fuel, oxygen and recycled flue gas and combust said fuel,oxygen and recycled flue gas to produce heat;

capturing heat produced by the oxy-fuel combustion;

using a portion of the heat to power a water splitter, therebygenerating hydrogen gas and oxygen gas; and

transferring the oxygen gas from the water splitter to the combustionchamber for use in said oxy-fuel combustion.

In another aspect, provided herein is a method for oxy-fuel combustion,comprising:

providing a system comprising a combustion chamber arranged and disposedto receive coal/water slurry, oxygen and recycled flue gas, wherein thechamber is arranged and disposed to receive said oxygen from an airseparator unit and/or external oxygen source, and/or a water splitter,and combust said coal/water slurry, oxygen and recycled flue gas toproduce heat and heated flue gas containing carbon dioxide; one or moreheat exchangers arranged and disposed to capture heat from said heatedflue gas and transfer a portion of the captured heat to a watersplitter; a water splitter using a 4-step hybrid copper-chlorinethermochemical cycle for the conversion of heat and/or electricity intohydrogen gas and oxygen gas, arranged and disposed to transfer theoxygen gas to the combustion chamber for use in said oxy-fuelcombustion;

combusting the coal/water slurry and oxygen to produce heat and heatedflue gas containing carbon dioxide;

capturing heat from heated flue gas and transferring captured said heatto the water splitter;

using a portion of the heat to power the water splitter using a 4-stephybrid copper-chlorine thermochemical cycle, thereby producing hydrogengas and oxygen gas;

transferring the oxygen gas to the combustion chamber for use in saidoxy-fuel combustion; and

reducing or eliminating the amount of oxygen that the combustion chamberrequires from an air separator unit and/or external oxygen supply, inproportion to the amount of oxygen received from the water splitter.

In yet another aspect, provided herein is a method for the reaction ofhydrogen produced by a water splitter and carbon dioxide obtained fromcombustion flue gas to form methanol and water.

In still aspect, provided herein is an oxy-fuel combustion system,comprising:

a combustion chamber arranged and disposed to receive fuel, oxygen andrecycled flue gas and combust said fuel, oxygen and recycled flue gas toproduce heat and heated flue gas containing carbon dioxide;

one or more heat exchangers arranged and disposed to capture heatproduced by the oxy-fuel combustion and transfer said heat to a watersplitter; and

a water splitter, for the conversion of heat and electricity intohydrogen gas and oxygen gas.

In another aspect, provided herein is an oxy-fuel combustion system,comprising:

a combustion chamber arranged and disposed to receive coal/water slurry,oxygen and recycled flue gas, wherein the chamber is arranged anddisposed to receive said oxygen from an air separator unit and/orexternal oxygen source, and/or a water splitter, and combust saidcoal/water slurry, oxygen and recycled flue gas to produce heat andheated flue gas containing carbon dioxide;

one or more heat exchangers arranged and disposed to capture heat fromsaid heated flue gas and transfer a portion of the captured heat to awater splitter; and

a water splitter using a 4-step hybrid copper-chlorine thermochemicalcycle for the conversion of heat and electricity into hydrogen gas andoxygen gas, arranged and disposed to transfer the oxygen gas to thecombustion chamber for use in said oxy-fuel combustion.

As described herein, these methods and systems of oxy-fuel combustionare recuperative.

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1 depicts the recuperative combustion process described herein,further comprising a methanol reactor, means for separating methanol andwater, and means for conversion of methanol to downstream products.

FIG. 2 depicts a recuperative combustion system Integrated with an 875MWTH (HHV) ISOTHERM PWR System (modified from Hong et al. 2008).

FIG. 3 depicts a flow diagram of a methanol synthesis plant (modifiedfrom Ushikoshi et al. 2000).

FIG. 4 depicts a recuperative combustion system compared to anarchetypal combustion system.

FIG. 5 depicts the steps of the four-step copper chlorine thermochemicalcycle (Wang et al. 2009).

FIG. 6 depicts an overview of the sulfur-iodine thermochemical cycle(Brown et al. 2003).

FIG. 7 depicts an archetypal convective recuperator (Bureau of EnergyEfficiency 2004).

FIG. 8 depicts an archetypal radiation/convective recuperator (Bureau ofEnergy Efficiency 2004).

FIG. 9 depicts an archetypal regenerator (Bureau of Energy Efficiency2004).

DETAILED DESCRIPTION OF THE INVENTION

In one aspect, provided herein is a method for oxy-fuel combustion,comprising: providing a system comprising a combustion chamber arrangedand disposed to receive fuel, oxygen and recycled flue gas and combustsaid fuel, oxygen and recycled flue gas to produce heat; capturing heatproduced by the oxy-fuel combustion; using a portion of the heat topower a water splitter, thereby generating hydrogen gas and oxygen gas;and transferring the oxygen gas from the water splitter to thecombustion chamber for use in said oxy-fuel combustion.

Combustion Chamber and Fuel

Oxy-fuel combustion for the production of electricity (FIG. 1-1) canoccur in a variety of combustion systems. Non-limiting examples ofcombustion systems include: circulating fluidized bed boiler, pulverizedcoal boiler, and combustors. Combustible materials, including, but notlimited to coal, coal/water slurry, petroleum products including oil,methane and natural gas, biomass, and plasma fuels (FIG. 1-3), mayundergo oxy-fuel combustion. The primary products of oxy-fuel combustioninclude heat and flue gas rich in carbon dioxide (FIG. 1-23). A portionof the flue gas is recycled, as described herein (FIG. 1-13).

In one embodiment of the method, the oxy-fuel comprises any combustiblematerial. In another embodiment, the oxy-fuel comprises anyhydrocarbon-based fuel. In yet another embodiment, the oxy-fuelcomprises coal/water slurry. In still another embodiment, the oxy-fuelcomprises oil.

Pressurized oxy-fuel combustion systems have the potential for betterperformance when compared to conventional atmospheric oxy-fuelcombustion power cycles such as the ITEA ISOTHERM® pressurized oxy-fuelsystem (Hong et al, 2008). For example, oxy-fuel combustion at highpressures may increase the burning rate of char and the heat transferrates in the convective sections of the heat transfer equipment.Further, because of the raised dew point and the corresponding availablelatent enthalpy in the raw flue gases, the pressurized oxy-fuel systemcan recover more thermal energy from the flue gases and eliminate thebleeding from the high-pressure and the low-pressure steam turbines.Consequently, the cycle efficiency for the pressurized oxy-fuel systemmay be superior to the atmospheric system. To operate this highcombustion pressure system, a high pressure deaerator and a flue gasacid condenser can be used. The acid condenser may be modified to workat a high pressure level with flue gas composition seen inoxy-combustion. (Hong et al. 2008).

Air Separator Unit and Water Splitter

In one embodiment of the method, the system comprises a water splitterarranged and disposed to provide oxygen to the combustion chamber foruse in oxy-fuel combustion.

In one embodiment of the method, the system further comprises an airseparator unit arranged and disposed to provide oxygen to the combustionchamber for use in oxy-fuel combustion.

The normal source of oxygen for oxy-fuel combustion is an air separatorunit (ASU) (FIG. 1-5). The selection and implementation of air separatorunits will be well known to those skilled in the art. Commonly, oxygenis produced in the ASU through a cryogenic distillation process.However, the production of oxygen from an ASU requires 15-20% of thegross power output of an industrial facility. The need for an ASU foroxygen generation may be reduced or eliminated through the use of awater splitter. Candidate water splitters include a hybridcopper-chlorine thermochemical Cycle (CuCl Cycle), sulfur-iodinethermochemical cycle, hybrid sulfur thermochemical cycle (HyS Cycle)and/or high temperature electrolysis (HTE). Other candidate watersplitters include volatile metal oxide thermochemical cycles(e.g.,zinc/zinc oxide, hybrid calcium), and non-volatile metal oxidecycles (e.g., iron oxide, cerium oxide) thermochemical cycles. Otherwater-splitting thermochemical cycles are available, and known to thoseskilled in the art. Water splitters use thermal or electrical energy inorder to convert water (FIG. 1-17) to hydrogen gas (FIG. 1-15) andoxygen gas (FIG. 1-11). In the recuperative combustion process describedherein, the water splitter is powered by excess process heat provided byoxy-fuel combustion (FIG. 1-7), and by electricity (FIG. 1-17). Oxygenproduced by a thermochemical cycle or HTE offsets, reduces or eliminatesthe need for an ASU, or external oxygen source, for oxy-fuel combustion,thereby increasing combustion efficiency. The thermochemical cycle orHTE unit employed in the system also produces hydrogen as a product ofwater splitting. The hydrogen may be used for a variety of purposes,including reaction with the carbon dioxide produced by oxy-fuelcombustion to yield methanol and water (FIG. 1-25).

In another embodiment of the method, the system further comprises an airseparator unit, wherein the combustion chamber is arranged and disposedto receive oxygen from the air separator unit and/or the water splitter.

Use of a water splitter to produce oxygen results in a reduction, orelimination, of the need for an air separation unit (ASU), or otheroxygen source, to supply oxygen for oxy-fuel combustion. The scaling ofthe water splitter determines if an ASU output is either reduced oreliminated. The employment of a water splitter may significantly reduce,or eliminate, the cost of an ASU unit and, correspondingly, reduce thepower penalty resulting from the operation of an ASU. Capturedprocess/waste heat, some of which may be upgraded by chemical heatpumps, may be used to supply additional power needs by the watersplitter as well.

In one embodiment of the method, the combustion chamber is arranged anddisposed to receive oxygen from the air separator unit and/or externaloxygen supply, and/or the water splitter.

The selection and implementation of suitable water splitters, apparatusand procedures for use in the methods and systems described herein willbe well known to those skilled in the art.

In a pressurized combustor (Hong et al. 2008), the oxygen puritydelivered from the water splitter and/or ASU may be 95% (mol %) anddelivered at 200° C.; oxygen content in the recycled flue gas may beabout 3% (mol %). The oxygen delivery temperature to the combustor iscontrolled to prevent the acid condensation when mixed with the recycledflue gases. The flue gases contain acid gases, such as SO₃, SO₂,nitrogen oxides, and HCl produced during combustion. As shown in FIG. 2,the recycled flue gases, state 2-19, are mixed with the oxygen streamfrom the water splitter, state 2-13, which may be colder, depending onthe ASU or water splitter technology employed. To avoid corrosion due tothe condensation of these acid gases when mixed with the oxygen stream,the oxygen delivery temperature needs to be carefully controlled tonearly 200° C. The 200° C. delivery temperature target is achieved byusing a two-stage oxygen compressor with an intercooler. The mass flowrate of the oxygen stream is determined such that the raw flue gasesexiting the combustor have 3% oxygen on a molar basis (Hong et al. 2008)

In one embodiment of the method, the amount of oxygen required from theair separator unit and/or external oxygen supply is reduced oreliminated in proportion to the amount of oxygen provided by the watersplitter. The amount of oxygen that is required from the air separatorunit, and/or another oxygen source, and/or the water splitter, may bedetermined by methods taught herein (see, for example, the section ofAlgorithms and Formulae), or by methods known to those of skill in theart.

Thermochemical Cycles

In another embodiment of the method, the water splitter produceshydrogen gas and oxygen gas by means of a thermochemical cycle. In stillanother embodiment, the water splitter produces hydrogen and oxygen gasby means of a hybrid thermochemical/electrochemical cycle.

The selection and implementation of suitable thermochemical or hybridthermochemical/electrochemical cycles, apparatus and procedures for usein the methods and systems described herein will be well known to thoseskilled in the art.

Non-limiting examples of thermochemical cycles and hybridthermochemical/electrochemical cycles include sulfur cycles (e.g.,hybrid sulfur and sulfur-iodine thermochemical cycles), low temperaturecycles (e.g., hybrid copper-chlorine thermochemical cycle), volatilemetal oxide cycles (e.g.,zinc/zinc oxide, hybrid calcium), andnon-volatile metal oxide cycles (e.g., iron oxide, cerium oxide)thermochemical cycles. Other suitable thermochemical water-splittingcycles are available, and known to those skilled in the art.

In one embodiment, the thermochemical cycle is selected from: a hybridcopper-chlorine cycle; a sulfur-iodine cycle; and a hybrid sulfur cycle.

In one embodiment, the thermochemical cycle is a hybrid copper-chlorinecycle. In another embodiment, the copper-chlorine cycle is selectedfrom: a 3-step cycle, a 4-step cycle, and a 5-step cycle. In yet anotherembodiment, the hybrid copper-chlorine cycle is the 4-step cycle. Instill another embodiment, the 4-step cycle of the hybrid copper chlorinecycle may be represented by the following steps:

-   -   Step I: Cu(s)+2HCl(g)→2CuCl(molten)+H₂(g)    -   Step II: 4CuCl(s)→2Cu(s)+2CuCl₂(aq)+HCl(aq)    -   Step III:        CuCl₂(aq)+n_(f)H₂O(l)→CuOCuCl₂(s)+2HCl(g)+(n_(f)−1)H₂O(g)    -   Step IV: CuOCuCl₂(s)→2CuCl(molten)+0.5O₂(g)

In certain embodiments, n_(f) is 5-30.

Three-step, four-step and five-step versions of the copper-chlorine(Cu—Cl) hybrid electrochemical-thermochemical cycle are described inWang et al. (2009). While the five-step version requires less heat thanthe three-step version of the hybrid Cu—Cl cycle, it is more complexfrom equipment and process engineering standpoints. A hybrid Cu—Cl Cycleusing the four-step process is described by Chukwu et al. (2008). StepIII of the four-step process is an electrochemical step, requiring theinput of electricity. Chukwu et al. provides reaction-specificthermodynamic data and Aspen Plus heat/mass balance modeling of thefour-step hybrid Cu—Cl cycle.

The four-step process (FIG. 5) consists of three thermal reactions inwhich H₂, O₂ and HCl are generated, and an electrochemical step in whichCuCl is disproportionated to yield copper metal and CuCl₂. The oxygen isreleased from the copper oxychloride between 450° C. and 530° C., whichis the highest temperature limit for this cycle (Sattler 2010).

In one embodiment of the method, the water splitter operates at atemperature 450° C. for Step I; 30-80° C. for Step II; 375° C. for StepIII; and 530° C. for Step IV. another embodiment, the heat required toproduce 1 mole of O and 1 mole of H₂ is about 554.7 kJ/mol.

The hybrid Cu—Cl four-step cycle receives process heat in adequatesupply from oxy-fuel combustion through a heat exchanger system. Thehybrid Cu—Cl four-step cycle is sized to accommodate oxygen requirementsfor oxy-fuel combustion within the fossil fuel combustor or boiler. Thissizing is based on either a partial or complete replacement of the airseparator unit (ASU), depending on a number of pre-existing, technicaland economic factors unique to an industrial/power plant, or other siteof combustion. Hydrogen production rates for various downstream hydrogenuses are also an important aspect of sizing the four-step hybrid Cu—Clthermochemical cycle reactor.

In one embodiment of the method, the thermochemical cycle is asulfur-iodine (S—I) cycle. In another embodiment, the sulfur-iodinecycle may be represented by the following steps:

-   -   Step I: 2H₂O+SO₂+I₂→H₂SO₄+2HI    -   Step II: H₂SO4→H₂O+SO₂+½O₂    -   Step III: 2HI→H₂+I₂

Heat and mass balances associated with the S—I Cycle can be found inBrown et al. (2003). The gross heat needed per mole H₂ (and 0.5 mole O₂)is 674.9 kJ/mole H₂ with net heat requirements is 391.3 kJ/mol H₂ (fullHI gasification), 432.9 kJ/mol H₂ (no HI gasification), as compared with554.7 kJ/H₂ gross heat and 322.7 kJ/H₂ net heat required for the Cu—Clcycle (Wang et al. 2009). A major difference between the Cu—Cl and S—Ithermochemical cycles is the substantially lower temperatures that theCu—Cl cycle operates at relative to the S—I cycle, allowing almost 30%of the heat needed for the Cu—Cl process to come from low grade heat(i.e., heat at temperatures lower than 343 K).

The hybrid sulfur (HyS) cycle is a hybrid electrochemical—thermochemicalcycle (Bilgen 1988). It consists of splitting sulfuric acid into waterand sulfur trixode (endothermic), followed by further decomposition ofsulfur trioxide to sulfur dioxide (highly endothermic). As a final step,sulfur dioxide is electrochemically oxidized to sulfuric acid withconcomitant production of hydrogen. The steps of the HyS cycle are asfollows:

-   -   Step I: H₂SO₄→H₂O+SO₃ (>450° C.)    -   Step II: SO₃→SO₂+½O₂ (>850° C. with catalyst; 1150° C. without        catalyst)    -   Step III: 2H₂O+SO₂→H₂SO₄+H₂ (electrolysis, 80° C.)

Electrical power is required for the electrolysis, but theelectrochemical oxidation of SO₂ is far more efficient than theelectrolytic splitting of water. The overall efficiency of the processis calculated to be about 40%. Carbon-supported platinum electrodes areused for the SO₂ oxidation. Cells made from ceramics such as siliconcarbide, silicon nitrite, and cermets possess excellent resistance tocorrosion by sulfuric acid at ambient temperature and at low acidconcentration. Catalysts mainly based on iron oxide are available foraccelerating the reaction rate of the SO₃ reduction at “low” temperature(850° C.). The kinetics of the reaction are much faster if highertemperatures are available as in solar tower installations. Therefore,the use of catalysts might be reduced or even unnecessary if thesulfuric acid splitting is coupled to concentrated solar radiation. Ithas to be evaluated whether the higher temperatures are more efficientthan the catalyzed reaction on an annual basis. Reactors used inlaboratory tests have been made of glass or fused silica; solar reactorsare mostly constructed from ceramics such as silicon carbide, butgold-coated steel has also been used (Noglik et al. 2009; Sattler 2010).Like the S—I Cycle, the HyS cycle may be integrated with therecuperative combustion process described herein. However, the highgrade heat (>850 C with catalyst, or 1150 C without catalyst) limits theapplication for combustion facilities to those applications with veryhigh heat output (i.e., oxy-fuel combustion with high oxygencontent/lower recirculated flue gases, or high temperature industrialprocesses such as, for example, steel milling or glass production).

High Temperature Electrolysis

In one embodiment of the method, the water splitter produces hydrogengas and oxygen gas by means of high-temperature electrolysis.

The selection and implementation of suitable high temperatureelectrolysis apparatus and procedures for use in the methods and systemsdescribed herein will be well known to those skilled in the art.

Hydrogen and Oxygen can be produced via the classical electrolysis ofwater at low temperature or, alternatively, by using the different fuelcell technologies. These technologies are based on (i) proton-exchangemembrane fuel cells (PEMFCs) (referring to the solid polymericelectrolyte membrane), (ii) fuel cells using solid oxide protonconductors, and (iii) fuel cells with a solid oxide ion (O²⁻) conductor(SOFCs). In a fuel cell, electrical energy is generated by theexothermic oxidation of hydrogen. In the reverse electrolysis operationof such a cell, steam is reduced in an endothermic reaction usingelectrical energy.

The operating temperatures of fuels cells vary widely, from around80-120° C. for PEMFCs to 700-1000° C. for SOFCs. The free energyrequired for the reaction (ΔG) decreases with increasing temperaturewhereas the free enthalpy (ΔH) remains almost constant. Thisthermodynamic relation, in principle unfavorable for the fuel cell modeat high temperatures, explains the particular interest in performingelectrolysis at high temperatures. Since the SOFCs achieve competitive(chemical-to-electrical) energy conversion efficiencies despite the lessfavorable thermodynamic conditions, one can a priori expect thathigh-temperature electrolysis (HTE) cells achieve much higher(electrical to-chemical) energy conversion efficiencies (the term energyconversion efficiency for the HTE refers to the electrical-to-chemicalenergy conversion). Because ionic transfer numbers are close to one forboth cell types, the difference in cell voltage translates linearly tothe energy consumption for the reaction.

The primary motivation for HTE is the above-mentioned potential of areduced demand for electrical energy compared with electrolysis at lowtemperature. This may allow electrical-to-chemical energy conversionefficiencies even exceeding 100%, as already recognized in early work(Isenberg 1981). The free energy of the reaction ΔG decreases from ˜1.23eV (237 kj mol⁻¹) at ambient temperature to −0.95 eV (183 kj mol⁻¹) at900° C., while the free enthalpy term remains essentially unchanged(ΔH≈1.3 eV or 249 kj mol⁻¹ at 900° C.). Part of the energy required foran ideal (loss-free) HTE can thus be provided by heat. Increasing ohmicand/or reaction losses in a real HTE system increase the demand forelectrical energy and decrease the demand for an external heat supplyuntil, finally, the reaction becomes exothermic. Hence three modes ofoperation are distinguishable in HTE: thermoneutral, endothermic, andexothermic. HTE operates at thermal equilibrium (the thermoneutral mode)when the electrical energy input equals the enthalpy of the reaction. Inthat case, the entropy necessary for water splitting equals the heatgenerated by the loss reactions, and the energy conversion efficiency is100%. In the exothermic mode, on the other hand, the electric energyinput exceeds the ΔH term, which corresponds to efficiency below 100%.Finally, in the endothermic mode, the electric energy input remainsbelow the enthalpy term. Therefore, heat must be supplied to maintainthe cell temperature. This mode means that energy conversionefficiencies of the cell or the stacks are above 100%.

An HTE system can be operated with and without an external heat supply.This is different to low-temperature electrolyzers, which run in theexothermic mode, because the energy losses, which arise mainly from theelectrochemical reactions, exceed the small difference between ΔH and ΔGat low temperature. The availability of an external heat sourceinfluences the design of an HTE system.

Without a heat source, the goal is to approach the thermoneutral mode,that is, to limit the thermal losses to a value required to compensatefor the endothermic reaction. This leaves a wide margin for cellovervoltages and, therefore, for an increase in the current density or alowering of the temperature. Operating temperatures in the range600-700° C., known from the SOFC development, may therefore also beaccessible for electrolysis.

With an external heat source of high temperature, on the other hand, thegoal is to reduce the overvoltages as far as possible to allow for asignificant uptake of heat. This implies, at least with present celltechnology, operation under higher temperatures (800° C. or above) andlower electrode overvoltages (i.e., current densities somewhat lowerthan those achieved in thermoneutral operation).

The operation of SOFCs in electrolysis mode has been demonstrated inseveral research projects since 2004 (EIFER 2010). Cells of bothcommercial and research types and including the common designs weretested (electrolyte, hydrogen electrode, and metal substrate-supported).As for fuel cell operation, the hydrogen electrode-supported cellsshowed the highest performance owing to the low resistance of the thinelectrolyte layer. A high current density of −3.6 A cm⁻² at a cellvoltage of 1.48 V and a cell temperature of 950° C., for example, wasreached with such a cell at DTU-Risoe (Denmark) (Mogensen et al. 2006;Zahid et al. 2010).

Heat Exchangers

In one embodiment of the method, the system further comprises one ormore heat exchangers for the capture and transfer of heat from theoxy-fuel combustion to the water splitter.

Heat exchangers are required for transferring process heat from theheated flue gas to endothermic reactions within the water splittersection. For example, in the Cu—Cl Cycle, sufficient heat must betransferred from the flue gas coming from the boiler/combustor to thechlorination step (Step I), oxychlorination step (Step (III) anddecomposition step (Step IV) (FIG. 5).

Numerous types of heat exchangers are available for the adequatetransfer of heat to the water splitter. Specific choice of heatexchanger is dependent on a number of factors, including, but notlimited to: the nature or quality of the flue gases, the temperature ofprimary process flue gases, matching heat demand of the secondary (i.e.,water splitting) process with the heat supply from the primary process,matching timing of the heat supply for the primary process and the heatdemand in the secondary process, and placement of primary and secondaryheating equipment.

The selection and implementation of suitable heat exchanger apparatusand procedures for use in the methods and systems described herein willbe well known to those skilled in the art.

In one embodiment of the method, the system further comprises one ormore heat exchangers for the capture and transfer of heat from theheated flue gas to the water splitter. In another embodiment, said oneor more heat exchangers are selected from: a convective recuperator, aradiation/convective recuperator, a ceramic recuperator and aregenerator. In yet another embodiment, said one or more heat exchangersare selected from appropriately scaled heat exchangers which functionwithin specified heat ranges (i.e., heat ranges specified herein orknown to those of skill in the art). Other suitable heat exchangers areavailable, and known to those of skill in the art.

Convective Recuperator. In a recuperator, heat exchange takes placebetween the flue gases and the air through metallic or ceramic walls.Ducts or tubes carry the air or other gas to be heated; the other sidecontains the waste heat stream. A common configuration for recuperatorsis called the tube type or convective recuperator. As seen in the FIG.7, the hot gases are carried through a number of parallel small diametertubes, while the incoming air/gas to be heated enters a shellsurrounding the tubes and passes over the hot tubes one or more times ina direction normal to their axes. If the tubes are baffled to allow thegas to pass over them twice, the heat exchanger is termed a two-passrecuperator; if two baffles are used, a three-pass recuperator, etc.Although baffling increases both the cost of the exchanger and thepressure drop in the combustion air path, it increases the effectivenessof heat exchange. Shell and tube type recuperators are generally morecompact and have a higher effectiveness than radiation recuperators,because of the larger heat transfer area made possible through the useof multiple tubes and multiple passes of the gases.

Radiation/convective Recuperator. For maximum effectiveness of heattransfer, combinations of radiation and convective designs are used,with the high-temperature radiation recuperator being first followed byconvection type. These are more expensive than simple metallic radiationrecuperators, but are less bulky. A Convective/radiative Hybridrecuperator is shown in FIG. 8.

Ceramic Recuperator. The principal limitation on the heat recovery ofmetal recuperators is the reduced life of the liner at inlettemperatures exceeding 1100° C. In order to overcome the temperaturelimitations of metal recuperators, ceramic tube recuperators have beendeveloped whose materials allow operation on the gas side to 1550° C.and on the preheated air side to 815° C. on a more or less practicalbasis. Early ceramic recuperators were built of tile and joined withfurnace cement, and thermal cycling caused cracking of joints and rapiddeterioration of the tubes. Later developments introduced various kindsof short silicon carbide tubes which can be joined by flexible sealslocated in the air headers. Earlier designs had experienced leakagerates from 8 to 60 percent. The new designs are reported to last twoyears with air preheat temperatures as high as 700° C., with much lowerleakage rates.

Regenerator. Regenerators are rechargeable storage batteries for heat. Aregenerator (FIG. 9) is an insulated container filled with metal orceramic shapes that can absorb and store relatively large amounts ofthermal energy. During the operating cycle, process exhaust gases flowthrough the regenerator, heating the storage medium. After a while, themedium becomes fully heated (charged). The exhaust flow is shut off andcold combustion air enters the unit. As it passes through, the airextracts heat from the storage medium, increasing in temperature beforeit enters the burners. Eventually, the heat stored in the medium isdrawn down to the point where the regenerator requires recharging. Atthat point, the combustion air flow is shut off and the exhaust gasesreturn to the unit. This cycle repeats as long as the process continuesto operate. For a continuous operation, at least two regenerators andtheir associated burners are required. One regenerator provides energyto the combustion air, while the other recharges. An alternate design ofregenerator uses a continuously rotating wheel containing metal orceramic matrix. The flue gases and combustion air pass through differentparts of the wheel during its rotation to receive heat from flue gasesand release heat to the combustion air. Regenerators may be preferablefor large capacities and have been very widely used in glass and steelmelting furnaces. Important relations exist between the size of theregenerator, time between reversals, thickness of brick, conductivity ofbrick and heat storage ratio of the brick. In a regenerator, the timebetween the reversals is an important aspect. Long periods would meanhigher thermal storage and hence higher cost. Also long periods ofreversal result in lower average temperature of preheat and consequentlyreduce fuel economy.

In one aspect, provided herein is a method for oxy-fuel combustion,comprising:

providing a system comprising a combustion chamber arranged and disposedto receive coal/water slurry, oxygen and recycle flue gas, wherein thechamber is arranged and disposed to receive said oxygen from an airseparator unit and/or a water splitter, and combust said coal/waterslurry, oxygen and recycled flue gas to produce heat and heated flue gascontaining carbon dioxide; one or more heat exchangers arranged anddisposed to capture heat from said heated flue gas and transfer aportion of the captured heat to a water splitter; a water splitter usinga 4-step hybrid copper-chlorine thermochemical cycle for the conversionof heat and/or electricity into hydrogen gas and oxygen gas, arrangedand disposed to transfer the oxygen gas to the combustion chamber foruse in said oxy-fuel combustion;

combusting coal/water slurry, oxygen and recycled flue gas to produceheat and heated flue gas containing carbon dioxide; capturing heat fromheated flue gas and transferring captured heat to the water splitter;Using a portion of the heat to power the water splitter using a 4-stephybrid copper-chlorine thermochemical cycle, thereby producing hydrogengas and oxygen gas; transferring the oxygen gas to the combustionchamber for use in said oxy-fuel combustion; reducing or eliminating theamount of oxygen that the combustion chamber requires from an airseparator unit and/or external oxygen supply, in proportion to theamount of oxygen received from the water splitter.

Efficiency

In one embodiment of the method, the amount of oxygen that thecombustion chamber requires from an air separator unit and/or externaloxygen supply, is reduced or eliminated in proportion to the amount ofoxygen received from the water splitter results in increased efficiency.

In one embodiment of the method, the increased efficiency is measured interms of increased gross power output of the combustion process. In oneembodiment, the gross power output of the combustion process isincreased by 1-20%. In another embodiment, the gross power output of thecombustion process is increased by 1-10%. In yet another embodiment, thegross power output of the combustion process is increased by 10-20%. Instill another embodiment, the gross power output of the combustionprocess is increased by 15-20%.

Primary and Secondary Products

In one embodiment of the method, the oxy-fuel combustion productsfurther comprise heated flue gas containing carbon dioxide.

Oxy-fuel combustion yields flue gases consisting of predominantly carbondioxide (FIG. 1-23) and condensable water, whereas conventionalair-fired combustion flue gases are nitrogen-rich with only about 15%(by volume) of carbon dioxide (Hu et al. 2008; IEA, 2008). The highcarbon dioxide concentration (up to 95%) and the significantly lowernitrogen concentration in the oxy-fuel raw flue gases is a uniquefeature that lowers the energy and capital costs of oxy-fuel carbondioxide capture when compared to alternatives (Buhre et al 2005).Further, decreasing relative recycled flue gas input increasescombustion, and flue gas, temperatures, and may be a mode for increasingheat grade for secondary water splitting processes. For example,oxy-fuel combustion, using a pressurized coal combustor, occurs at 1400to 1600° C.; however, stoichiometric combustion of coal in pure oxygenreaches up to 3500° C. Optimization of oxygen/recycled gas content withrespect to high grade heat production and oxygen/hydrogen yieldsrequires additional research.

Oxy-fuel combustion involves the burning of fuel in an oxygen-rich,nitrogen-lean and carbon dioxide-rich environment, which is achieved byfeeding the combustor or boiler with an oxygen-rich stream and recycledflue gases. Oxy-fuel combustion produces a flue gas stream containingmostly CO₂, which can be directly purified and compressed for conversionto useful materials or for carbon sequestration purposes. In the processshown in FIG. 1, the CO₂ concentration in the flue gas is greatlyincreased by using a mixture of recirculated flue gas and pure oxygeninstead of air for firing coal. Recirculation of flue gas is necessaryto provide sufficient mass flow of flue gas for cooling the flame andalso heat capacity and flue gas velocity for convective heat transfer inthe boiler. In the oxy-fuel process, CO₂ purity is mainly influenced by(a) where the flue gas is recycled in the process (the cleaning that hasbeen done up to this point), (b) the sealing of boiler and othercomponents to prevent air ingress, (c) the purity of the oxygen from theAir Separation Unit (ASU) (or alternative oxygen source), (d) theperformance of all air quality control system equipment (e.g., SCR, FGD,and ESP), and (e) additional CO₂ purification during/after compression(Wu et al. 2009). Oxy-fuel combustion involves using a mixture ofrecirculated flue gas and pure oxygen, instead of air, for firing thefuel. Oxy-fuel combustion may take place at 1400 to 1600° C. Thestoichiometric combustion of coal in pure oxygen may reach up to 3500°C. While oxy-fuel combustion is close to stoichiometric, lowercombustion temperature are achieved by using the appropriate amount ofthe recycled flue gases (Hong et al 2008). Further, certain additionalefficiencies are realized when oxy-fuel combustion takes place underpressurized conditions (Hong et al. 2008).

The primary outputs of the recuperative combustion process include thefollowing:

CO₂ (and small amounts of CO) in the flue gas, following post combustiontreatment;

-   H₂ and O₂ from the water splitter (i.e., Cu—Cl, S—I, or HyS    thermochemical cycles) or HTE. The secondary outputs from the    downstream portions of the process include: methanol from the    methanol reactor (FIG. 1-25), dimethyl ether (DME), olefins, and    gasoline (mainly C5-C9) from methanol to gasoline conversion (FIG.    1-35), and water, as a by-product of methanol and other hydrocarbon    production. This water may be treated, as necessary, and recycled to    the water splitter (FIGS. 1-31 and 1-41). The secondary outputs of    the recuperative combustion process described herein are valuable    products on the global market (FIG. 1-39).

In one embodiment of the method, the hydrogen from the water splitter isused directly or indirectly in a subsequent process. In anotherembodiment of the method, the hydrogen from the water splitter is useddirectly in a subsequent process.

In another embodiment of the method, the hydrogen from the watersplitter and the carbon dioxide from the combustion flue gas are reactedto form methanol and water.

In one aspect, provided herein is a method for the reaction of hydrogenproduced by a water splitter and carbon dioxide obtained from combustionflue gas to form methanol and water. In one embodiment, the carbondioxide is purified and compressed prior to reacting with hydrogen. Inanother embodiment, the hydrogen and carbon dioxide are both produced bythe recuperative combustion system. In yet another embodiment of themethod, the products of the reaction further comprise waste heat.

Using carbon dioxide or carbon monoxide for downstream conversion tohydrocarbons requires carbon dioxide separation and removal of theimpurities in the high-concentration carbon dioxide flue gases resultingfrom oxy-fuel combustion. Carbon dioxide purification involves theremoval of contaminants from the flue gas, including nitrogen oxides,sulfur oxides, and mercury, generally under pressurized conditions.While numerous strategies for removal of these contaminants alreadyexist at modern fossil fuel power plants, including flue gasdesulfurization (FGD) for sulfur oxides, selective catalytic reduction(SCR) for nitrogen oxides compounds and activated carbon or sorbents formercury, these are capital-intensive technologies that do not result ina comprehensive removal of impurities from, and separation of, carbondioxide from flue gas. A host of developing technologies for increasingcarbon dioxide purity are currently under development (e.g., see Whiteand Fogash 2009; Hong et al., 2008; and Shah 2006). Specifictechnologies employed for contaminant removal in the carbon dioxidestream depends on the end use of carbon dioxide (i.e., for methanolproduction, enhanced oil recovery (EOR) or sequestration).

The high-concentration carbon dioxide flue gas that is produced by theoxy-fuel combustion process may be hydrogenated in a fluidized bedreactor with hydrogen gas at 200-300° C. at a pressure of 50-100 bar ina catalytically-mediated reaction (heterogeneous catalyst includes, butis not limited to: Cu/ZnO/ZrO₂/Al₂O₃/SiO₂), yielding methanol, water andsubstantial heat (FIGS. 1-25, 1-21 and 1-31). The following threereactions are controlled in the methanol reactor to maximize methanolsynthesis (Hirotani et al. 1998):

CO+2H₂→CH₃OH (ΔH=−90.6 kJ/mol),

CO₂+3H₂→CH₃OH+H₂O (ΔH=−49.4 kJ/mol), and

CO₂+H₂→CO+H₂O (ΔH=+41.2 kJ/mol).

The selection and implementation of suitable apparatus and proceduresfor the production of methanol for use in the methods and systemsdescribed herein will be well known to those skilled in the art.

The methanol reactor may be of several types, including a modified testmethanol synthesis reactor from Ushikoshi et al. (2000). FIG. 3 shows aflow diagram of the test plant, which is designed with facilities forrecycling unreacted gases. The gases (mixture of CO₂, CO and H₂)supplied from the CO₂ purification and compression unit (mainly CO₂) andthe water splitter (H₂) (FIGS. 3-1, 3-2, 3-3, 3-4) are compressed (FIGS.3-5, 3-6, 3-7) along with recycled gases (FIG. 3-23), and then fed intothe reaction tube (FIGS. 3-8, 3-13) through a pre-heater (FIG. 3-9). Thereactor (FIG. 3-14) is surrounded by a heat exchanger divided into fourparts to facilitate isothermal operation, and capture of waste heat. Thetemperature profile along the bed is measured by means of eightthermocouples situated at the central axis of the reactor. Thetemperature difference along the reactor is less than 2° K. The pressureis controlled within 0.1 MPa by changing the total flow rate of themake-up gas, in which the H₂/CO/CO₂ ratio is adjusted with the flowcontrollers.

The flow rate of the inlet gas to the reactor is controlled by the flowcontroller placed just after the recycle gas compressor. Reactionproducts are cooled down (FIGS. 3-18, 3-19), and then the mixture ofmethanol and water is separated at the gas-liquid separator (FIGS. 3-20,3-21) from unreacted gases. Unreacted gases and gaseous products, suchas CO, methane and so on, excluding small amounts of purge gas, arerecycled back to the reactor (FIGS. 3-22, 3-23)). The mixture ofmethanol and water is subjected to separation (see below) then stored ina container (FIGS. 3-25, 3-26, 3-27). Control of temperature in themethanol reactor for pre-reaction reduction of the catalyst (in thepresence of heat and hydrogen and nitrogen gases) as well as duringmethanol synthesis is controlled by a preheater, oil heater and oilcooler system with (FIGS. 3-9, 3-10, 3-11, 3-12, 3-13, 3-14, 3-15, 3-16and 3-17).

The make-up gas, the inlet and outlet gases of the reactor and therecycle gas are analyzed with an on-line gas chromatograph. Gaschromatography is employed for analysis of the reaction products; H₂, COand CO₂ are analyzed by thermal-conductivity detector; methanol,dimethyl ether, methyl formate and hydrocarbons are analyzed by theflame ionization detector. Excess heat produced by exothermic methanolreaction (FIG. 3-30) may be recuperated and recirculated for use in thewater splitter (FIG. 3-32) (through heat upgrading using chemical heatpumps) (FIG. 3-31).

The methanol and water are separated through a solvent dehydrationprocess (e.g., by energetically attractive pervaporation using ahydrophilic ZeoSep A membrane or equivalent hydrophobic membrane forconcentration of organics, or a distillation column). Final selection ofappropriate pervaporation membrane or distillation column design shalldepend on optimized methanol and water concentrations in themethanol/water mixture.

The methanol may be used as is, or used as an industrial feedstock(e.g., for conversion to dimethyl ether, olefins, gasoline and aspectrum of other industrial applications) using a variety of industrialprocesses such as the Mobil Methanol to Gasoline process. The byproductwater is filtered and re-used in the in the water splitter.

Chemical Heat Pumps

Chemical Heat Pumps (CHPs) are systems that use coupled exothermic andendothermic reactors to store thermal energy and transform it to anothertemperature, including waste heat whose thermal energy at lowtemperatures can be upgraded to higher temperatures (Naterer 2008). CHPsmay be useful in the conversion of waste heat captured from exothermicportions of the Water Splitter Section or downstream Methanol Reactorfor Methanol to Gasoline Sections and upgraded for use inheat-intensive, endothermic portions of the Water Splitter Section.Gainful use of CHPs result in significant increase in hydrogen andoxygen production efficiency, and overall oxy-fuel combustion systemefficiencies.

The selection and implementation of suitable chemical heat pumpapparatus and procedures for use in the methods and systems describedherein will be well known to those skilled in the art.

Two specific solid-gas CHPs, namely salt/ammonia and MgO/water systems,are particularly useful in application to waste heat upgrading forthermochemical hydrogen production, especially when configured in series(Naterer 2008). These are described below.

A salt/ammonia chemical heat pump consists of salts that are able toabsorb/desorb ammonia vapor at different operating temperatures. Theammonia vapor pressure is a function of temperature for two differentsalts, designated by LTS (low-temperature salt) and HTS(high-temperature salt). The desorption reaction is endothermic. Heatmust be supplied to the gas/solid reactor to release ammonia vapor fromthe LTS. When this ammonia vapor flows to the HTS, it is absorbed andheat is released in an exothermic reaction. The pair of salts isMnSO₄/NH₃ (LTS) and NiCl₂/NH₃ (HTS). The operation of the chemical heatpump is described below:

MnSO₄.6NH₃ +Q _(waste)

4MnSO₄.2NH₃+4NH₃; ΔH=+57.6 kJ/mol (NH₃)

NiCl₂.2NH₃+4NH₃

NiCl₂.6NH₃ +Q _(out); ΔH=−55.3 kJ/mol (NH₃)

An integrated closed cycle of a salt/ammonia chemical heat pump waspresented and analyzed by Spoelstra et al. (2002). In their analysis,5000 kW of low-temperature heat at 140° C. was upgraded to 2051 kW ofhigh-temperature heat at 240° C. Shell-and-tube reactors were used withfinned tubes to achieve this operating capacity. Each reactor vessel wasabout 6 m in height, with a diameter of 2-3 m. The total weight of onevessel was about 50 tons, including the salt and heat exchanger tubes. Avery high coefficient of performance for this CHP was reported by theauthors (COP=97), since electrical power is only required to pump aroundliquid streams. This chemical heat pump could be used as a “bottomingcycle” to upgrade waste heat to an intermediate stage, before anotherCHP upgrades further to higher temperatures. It is anticipated thatequipment performance and reaction kinetics would become unfavorable ifa single CHP attempts to operate over an excessively large temperaturerange. Therefore, a magnesium oxide (MgO)/vapor chemical heat pump,which operates at temperatures above the salt/ammonia CHP may be usefulin increasing the grade of heat further.

The MgO/Vapor CHP is described by the following chemical reaction:

MgO(s)+H₂O(g)

4Mg(OH)2; ΔH=_(—)−81.02 kJ/mol

The rightward reaction is exothermic MgO hydration. The kinetics of thereaction have been reported by Kato et al. (1996). The operationconsists of heat storage and heat supply modes, with solid products fromeach reactor supplied as solid feed to the other. Magnesium hydroxide(Mg(OH)₂) is initially charged into a gas/solid reactor. Heat is added,after which solid MgO and water vapor are formed. The heat ofcondensation is recovered from the steam and the resulting water isstored as a liquid. In the heat supply mode, the stored water is thenvaporized by another separate heat input. The vapor is supplied to anexothermic solid/gas reactor for hydration of MgO. Scientificfeasibility of the MgO/vapor chemical heat pump has been demonstrated byKato et al. (1996). A lab-scale demonstration was performed by theauthors, with an average heat output rate of 349 W per kg of Mg(OH)₂solid feed. The experimental apparatus consisted of an evaporator,gas/solid reactor, heating supply with an electric furnace, condenser,water trap and vacuum pump. Future research and development are stillneeded to scale up this system to larger heating capacities (Naterer2008).

Using a sequence of chemical heat pumps in series, a conceptualframework of coupled CHPs and a Cu—Cl thermochemical cycle may beadvantageous. Specifically, an exothermic step within the Cu—Cl cycle(or Methanol Reactor Section) could supply heat into the MgO/vapor CHP,to be subsequently upgraded to a higher temperature that is then used bythe endothermic hydrolysis step in the Cu—Cl cycle. The lowertemperature endothermic step of copper oxychloride decomposition couldthen be supplied separately from the salt/ammonia CHP. Input power isneeded to drive compressors in the CHPs. With existing heat recoverytechnology available in commercial systems, electricity generated fromwaste heat could be supplied directly to the CHPs, thereby potentiallymaking the CHPs and Cu—Cl cycle solely driven by process/waste heat fromindustrial or power plants (Naterer 2008).

Naterer (2008) presents a thermodynamic analysis of combined chemicalheat pumps for a thermochemical cycle of hydrogen production,demonstrating that low-grade waste heat can be upgraded to highertemperatures via salt/ammonia and MgO/vapor chemical heat pumps, whichrelease heat at successively higher temperatures through exothermicreactions. Using this new approach, waste heat industrial sources can betransformed to a useful energy supply for thermochemical hydrogen andoxygen production. Naterer (2008) further provides an example theapplication of salt/ammonia and MgO/vapor chemical heat pumps to theCu—Cl thermochemical cycle production of oxygen and hydrogen.

In one embodiment, thermal energy from exothermic processes of themethod is captured by one or more chemical heat pumps. In anotherembodiment, said thermal energy is transformed to another temperature.In yet another embodiment, said thermal energy is used in endothermicprocesses of the method. In certain embodiments, thermal energy from thewater splitter cycle and/or the methanol reactor is captured,transformed to another temperature, and used in endothermic processes ofthe method.

System

In one aspect, provided herein is an oxy-fuel combustion system,comprising:

a combustion chamber arranged and disposed to receive fuel, oxygen andrecycled flue gas and combust said fuel, oxygen and recycled flue gas toproduce heat and heated flue gas containing carbon dioxide; one or moreheat exchangers for capturing heat produced by the oxy-fuel combustionand transferring said heat to a water splitter; and a water splitter,for the conversion of heat and electricity into hydrogen gas and oxygengas.

In one embodiment of the system, the water splitter is arranged anddisposed to provide oxygen to the combustion chamber for use in oxy-fuelcombustion.

In one embodiment, the system further comprises an air separator unitarranged and disposed to provide oxygen to the combustion chamber foruse in oxy-fuel combustion. In another embodiment, the system furthercomprises an external oxygen supply.

In one embodiment, the system further comprises an air separator unit,wherein the combustion chamber is arranged and disposed to receiveoxygen from the air separator unit and/or external oxygen supply and/orthe water splitter.

In one embodiment of the system, one or more heat exchangers serve asthe means to capture the heat produced by the oxy-fuel combustion and totransfer said heat to the water splitter. In another embodiment, saidone or more heat exchangers captures heat from the heated flue gas. Inyet another embodiment, said one or more heat exchangers are selectedfrom: a convective recuperator, a radiation/convective recuperator, aceramic recuperator and a regenerator. In still another embodiment, saidone or more heat exchangers are selected from appropriately scaled heatexchangers which function within specified heat ranges (i.e., heatranges specified herein or known to those of skill in the art).

In one embodiment of the system, the amount of oxygen required from theair separator unit and/or external oxygen supply is reduced oreliminated in proportion to the amount of oxygen provided by the watersplitter. The amount of oxygen that is required from the air separatorunit, and/or another oxygen source, and/or the water splitter, may bedetermined by methods taught herein (see, for example, the section ofAlgorithms and Formulae), or by methods known to those of skill in theart.

In one embodiment of the system, the oxy-fuel combustion productsfurther comprise heated flue gas containing carbon dioxide.

In one embodiment of the system, the oxy-fuel comprises any combustiblematerial. In another embodiment, the oxy-fuel comprises anyhydrocarbon-based fuel. In yet another embodiment, the oxy-fuelcomprises coal/water slurry. In still another embodiment, the oxy-fuelcomprises oil.

In one embodiment of the system, the water splitter produces hydrogengas and oxygen gas by means of high-temperature electrolysis.

In one embodiment of the system, the water splitter produces hydrogengas and oxygen gas by means of a thermochemical cycle. In anotherembodiment, the thermochemical cycle is selected from: a hybridcopper-chlorine cycle; a sulfur-iodine cycle; and a hybrid sulfur cycle.In yet another embodiment, the thermochemical cycle is a sulfur-iodinecycle. In still another embodiment, the sulfur-iodine cycle may berepresented by the following steps:

-   -   Step I: 2H₂O+SO₂+I₂→H₂SO₄+2HI    -   Step II: H₂SO4→H₂O+SO₂+½O₂    -   Step III: 2HI→H₂→I₂

In one embodiment of the system, the thermochemical cycle is a hybridcopper-chlorine cycle. In another embodiment, the hybrid copper-chlorinecycle is selected from: a 3-step cycle, a 4-step cycle, and a 5-stepcycle. In yet another embodiment, the hybrid copper-chlorine cycle isthe 4-step cycle. In still another embodiment, the 4-step hybrid copperchlorine cycle may be represented by the following steps:

-   -   Step I: Cu(s)+2HCl(g)→2CuCl(molten)+H₂(g)    -   Step II: 4CuCl(s)→2Cu(s)+2CuCl₂(aq)+HCl(aq)    -   Step III:        CuCl₂(aq)+n_(f)H₂O(l)→CuOCuCl₂(s)+2HCl(g)+(n_(f)−1)H₂O(g)    -   Step IV: CuOCuCl₂(s)→2CuCl(molten)+0.5O₂(g)

In certain embodiments, n_(f) is 5-30.

In one embodiment of the system, the water splitter operates at atemperature 450° C. for Step I; 30-80° C. for Step II; 375° C. for StepIII; and 530° C. for Step IV. In another embodiment, the heat requiredto produce 1 mole of O and 1 mole of H₂ is about 554.7 kJ/mol.

In one embodiment of the system, the hydrogen from the water splitter isused directly or indirectly in a subsequent process. In anotherembodiment of the system, the hydrogen from the water splitter is useddirectly in a subsequent process.

In one embodiment of the system, the hydrogen from the water splitterand the carbon dioxide from the combustion flue gas are reacted to formmethanol and water.

In one aspect, provided herein is an oxy-fuel combustion system,comprising:

a combustion chamber arranged and disposed to receive coal/water slurry,oxygen and recycled flue gas, wherein the chamber is arranged anddisposed to receive said oxygen from an air separator unit and/orexternal oxygen source, and/or a water splitter, and combust saidcoal/water slurry, oxygen and recycled flue gas to produce heat andheated flue gas containing carbon dioxide; one or more heat exchangersarranged and disposed to capture heat from said heated flue gas andtransfer a portion of the captured heat to a water splitter; a watersplitter using a 4-step hybrid copper-chlorine thermochemical cycle forthe conversion of heat and electricity into hydrogen gas and oxygen gas,arranged and disposed to transfer the oxygen gas to the combustionchamber for use in said oxy-fuel combustion.

Integration of the unit processes of the system described herein will bewell known to those of skill in the art. An example of an integratedsystem is depicted in FIG. 2 and described in Table 1.

Module

In one aspect, provided herein is a method for converting a non-oxy-fuelcombustion system into a recuperative oxy-fuel combustion system, saidmethod comprising:

Providing a system comprising an air separator unit and/or externaloxygen source; one or more heat exchangers; a thermochemical and/orelectrochemical water splitter; and a flue gas converter; and convertinga non-oxy-fuel combustion system into a recuperative oxy-fuel combustionsystem.

In one embodiment of the method, the non-oxy-fuel combustion systemcomprises a combustion chamber.

In one embodiment of the method, the water splitter is arranged anddisposed to provide oxygen to the combustion chamber for use in oxy-fuelcombustion.

Algorithms and Formulae

The following algorithms and formulae pertain to system scaling.Underlying chemical formulae pertaining to water splitters may be foundin the Thermochemical Cycles section. System scaling is based on themaximum mass flow rate of the fuel (e.g., coal) used in the boiler orcombustor, which in turn is based on the gross cumulative power producedby the entire power plant (MWg), not considering the power penaltyincurred by equipment at the plant. The following are specific scalingconsiderations for the combustion system described herein. Specifically,system scaling is based on the following method (modified from Rubin etal. 2007):

-   -   i. Calculate maximum fuel flow rate;    -   ii. Calculate oxygen requirement to meet this fuel flow rate;    -   iii. Considerations for sizing the water splitter to meet the        oxygen flow rate    -   iv. Calculate CO₂ flow rate from combustion    -   v. Calculate hydrogen flow rate for converting carbon dioxide to        methanol; and    -   vi. Considerations for sizing the methanol to gasoline section.        These considerations are described below.        i. Calculate Maximum Fuel Flow Rate

Fuel flow rate is calculated based on a plant's gross cumulative power(MWg), heat rate and fuel properties (heating value). The relationshipin the equation below can be used to determine the fuel flow raterequired to generate the desired (or actual) gross power, given the fuelproperties and gross heat rate.

$M_{coal} = \frac{{MW}_{g} \times {HR}_{steam}}{2 \times \eta_{boiler} \times {HHV}_{coal}}$

wherein,

-   M_(coal)=mass flow rate of fuel (ton/hr)-   MW_(g)=gross cumulative power produced by the entire power plant;    this does not consider power used by equipment in the power plant    (MW)-   HR_(steam)=heat rate of the steam cycle, which excludes the effects    of the boiler efficiency (Btu/kWh)-   η_(boiler)=boiler efficiency (fraction)-   HHV_(Coal)=higher heating value of the coal on a wet basis (Btu/lb)    ii. Calculate Oxygen Requirement to Meet this Fuel Flow Rate

The maximum rate of oxygen produced by the water splitter is determinedas follows:

-   -   a) Calculate the stoichiometric O₂ requirement based on the fuel        flow rate, fuel composition, and emission factors for incomplete        combustion reactants;    -   b) Calculate the total O₂ requirement based on the excess oxygen        specified (approximately 3-5% excess); and    -   c) Calculate the total oxygen product (i.e., oxidant) flow rate        based on the oxygen purity (≧95%) and total O₂ requirement.

This oxygen flow should replace the oxygen that would have been producedby the Air Separation Unit (ASU) in a normal oxy-fuel system.

iii. Considerations for Sizing the Water Splitter to Meet the OxygenFlow Rate

The water splitter (i.e., Cu—Cl or S—I or HyS thermochemical cycles; orHTE) is sized to produce the maximum oxygen flow rate, thereby replacingthe ASU. Sizing the water splitter such that oxygen flow rate is lowerthan the maximum flow rate for the facility may result in the need tosupplement oxygen production with another source (e.g., an ASU).

The water splitter is sized based on stoichiometric oxygen output fromparticular water splitter reactions to meet the O₂ flow rate determinedabove.

iv. Calculate CO₂ Flow Rate from Combustion

The carbon dioxide mass flow rate, m_(FG), can be derived based on thefollowing equation:

m _(CO2) =[m _(COAL)(1−% ash)+m _(RFG) +m _(O2) −m_(IMPURITIES)](CO2_(CAPTURE))(CO2_(PURITY))

Wherein:

-   m_(FG)=flue gas mass flow rate-   % ash=fuel ash content, mass fraction-   m_(RFG)=recycle flue gas mass flow rate-   m_(Impurities)=mass of impurities removed during CO2 purification    (i.e., NOx, SOx and Hg)-   CO_(2Capture)=capture efficiency of CO2 in the flue gas-   CO_(2pUrity)=carbon dioxide purity requirement (generally≧95%)

Zhou et al. (2010) found that oxy-fuel combustion in a conventionalutility boiler had an ideal flue gas recycle (FGR) ratio generallyaround 0.7-0.75; whereas Hong et al. (2008) found a flue gas recycleratio in a pressurized coal combustor to be about 0.78. Flue gas ratiosdepends on boiler exit O₂ and fuel properties such that flue gas recycleratio is a linear function of the boiler exit O₂ and increases slightlywith air-to-fuel ratio.

v. Calculate Hydrogen Flow Rate for Converting Carbon Dioxide toMethanol

The hydrogen flow rate is determined stoichiometrically based on thefollowing equation:

CO₂+3H₂→CH₃OH+2H₂O

Thus, the flow rate of hydrogen is three times the flow rate of carbondioxide on a molar basis. This likely will be offset somewhat by theoverall efficiency of carbon dioxide to methanol conversion, which isdependent on specific methanol reactor employed. The methanol reactorshall be sized to accommodate maximum carbon dioxide flow, based onmaximum fuel flow rates.

vi. Considerations for Sizing the Methanol to Gasoline Section

The methanol to gasoline reactor system shall be sized to accommodatemaximum methanol flow, based, in turn, on maximum fuel and carbondioxide flow rates, respectively. Methanol to gasoline reactor sizing isbased on stoichiometric considerations of the following overallreactions:

The dimethyl ether product is then further dehydrated over a zeolitecatalyst (preferred zeolites may include ZSM-5, ZSM-11, ZSM-12, ZSM-35,and ZSM-48) to give a gasoline with 80% C5+ hydrocarbon products.Conversion efficiencies for methanol to gasoline conversion shall alsobe considered when sizing the reaction system.

Like the Methanol Section, the Methanol to Gasoline Section producessubstantial amounts of heat which may be recuperated and transferredthrough chemical heat pumps and/or other heat exchangers to the WaterSplitter Section to power endothermic reactions. Therefore, theseSections should be sized and positioned in a manner which facilitateswaste heat recuperation and transfer.

REFERENCES

The following publications are incorporated herein by reference:

-   1. ANSI/ISA-77.13.01-1999-Fossil Fuel Power Plant Steam Turbine    Bypass System. IN: Z. L. Wang*, G. F. Naterer, K. S. Gabriel, R.    Gravelsins, V. N. Daggupati. 2010. Comparison of sulfur-iodine and    copper-chlorine thermochemical hydrogen production cycles.    International Journal of Hydrogen Energy. 35 (2010). Pp. 4820-4830.-   2. ANSI/ISA-77.44.01-2007-Fossil Fuel Power Plant—Steam Temperature    Controls. IN: Z. L. Wang*, G. F. Naterer, K. S. Gabriel, R.    Gravelsins, V. N. Daggupati. 2010. Comparison of sulfur-iodine and    copper-chlorine thermochemical hydrogen production cycles.    International Journal of Hydrogen Energy. 35 (2010). Pp. 4820-4830.-   3. Bilgen, E. 1988. Solar Hydrogen Production by Hybrid    Thermochemicl Processes. Solar Energy, 41(2). Pp. 199-206. IN:    Sattler, C., 2010. Thermochemical Cycles. IN: Hydrogen Energy.    Edited by Detlef Stolten. Wiley-VC Verlag GmbH & Co. KGaA, Weinheim.    ISBN: 978-3-527-32711-9. pp. 189-206.-   4. Brown, L. C., Lentsch, R. D., Besenbruch, G. E., Schultz, K. R.,    and J. E. Funk. 2003. Alternative Flowsheets for the Sulfur-Iodine    Thermochemical Hydrogen Cycle. GA-A24266. General Atomics. San    Diego, Calif. 19 pp.-   5. Buhre B J P., Elliott L K., Sheng C D., Gupta R P., Wall T F.    Oxy-fuel combustion technology for coal-fired power generation.    Progress in Energy and Combustion Science, 2005. 31(4): p. 283-307.-   6. Bureau of Energy Efficiency. 2004. Waste Heat Recovery (Chapter    8). See: http://www.em-ea.org/Guide%20Books/book    2/2.8%20Waste%20Heat%20Recovery.pdf. Accessed July 2010. 18 pp.-   7. Chukwu, C. C, Naterer, G. F. and M. A. Rosen. 2008 PROCESS    SIMULATION OF NUCLEAR-BASED THERMOCHEMICAL HYDROGEN PRODUCTION WITH    A COPPER-CHLORINE CYCLE. University of Ontario Institute of    Technology. 9 pp.-   8. Dokiya D, Kotera Y. Hybrid cycle with electrolysis using a Cu—Cl    system. International Journal of Hydrogen Energy 1976;1: 117-21.-   9. Drbal L F, Boston P G, Westra K L, Black, Veatch. Power plant    engineering. Springer; 1996. ISBN 0412064014, 9780412064012.    Chapters 7-9, pp. 185-286. IN: Z. L. Wang*, G. F. Naterer, K. S.    Gabriel, R. Gravelsins, V. N. Daggupati. 2010. Comparison of    sulfur-iodine and copper-chlorine thermochemical hydrogen production    cycles. International Journal of Hydrogen Energy. 35 (2010). Pp.    4820-4830.-   10. European Institute for Energy Research (EIFER). 2010. Hi₂H₂    Product (“Highly Efficient, High Temperature, Hydrogen Production by    Water Electrolysis”). Within the European Framework Program 6, the    Deutsches Zentrum fur Luft-und Raumfahrt (DLR), the Danish Technical    University (DTU) and the Risoe National Laboratory (DTU-Risoe), and    the Swiss Federal Laboratories for Matierals Testing and research    (EMPA), http://hi2h2.com (accessed 15 Feb. 2010). IN: Mohsine, Z.,    Schefold, J., and A. Brisse. 2010. High-Temperature Electrolysis    Using Planar Solid Oxide Fuel Cell Technology: A Review. IN:    Hydrogen Energy. Edited by Detlef Stolten. Wiley-VC Verlag GmbH &    Co. KGaA, Weinheim. ISBN: 978-3-527-32711-9. pp. 227-242.-   11. Giaconia, A. , Grena, R., Lanchi, M., Liberatore, R., and P.    Tarquini. 2007. Hydrogen/methanol production by sulfur-iodine    thermochemical cycle powered by combined solar/fossil energy.    International Journal of Hydrogen Energy. Volume 32. pp. 469-481-   12. Gorensek, M. B. 2010. The Effect of Anolyte Product Acid    Concentration on Hybrid Sulfur Cycle Performance. Presented at World    Hydrogen Energy Conference. May 16-21, 2010. Essen Germany.-   13. Hirotani, K., Hitoshi, N., and K. Shoji. 1998. Optimum Catalytic    Reactor Design for Methanol Synthsis for TEC MRF-Z ® Reactor.    Catalysis Surveys from Japan 2 . Pp 99-106.-   14. Hong, J., Chaudhry, G., Brisson, J. G., Field, R., Gazzino, M.,    and A. F. Ghoniem. 2008. Analysis of Oxy-Fuel Combustion Power Cycle    Utilizing a Pressurized Coal Combustor. MIT Energy Institute.    Cambridge, Mass. 38 pp.-   15. Hu Y., Naito S., Kobayashi N., Hasatani M. CO2, NOx and SO2    emissions from the combustion of coal with high oxygen concentration    gases. Fuel, 2000. 79(15): p. 1925-1932. IN: Hong et al. (2008).-   16. IEA Greenhouse Gas R&D Programme Improvement in power generation    with postcombustion capture of CO2, report no. PH4/33. Cheltenham,    UK: IEA Greenhouse Gas R&D Programme, 2004. IN: Hong et al. (2008).-   17. Intergovernmental Panel on Climate Change (IPCC). IPCC special    report on carbon dioxide capture and storage. Cambridge, UK: IPCC,    2005.-   18. Isenberg, A. O. 1981. Solid State Ionics. 3/4 . p. 431. IN:    Mohsine, Z., Schefold, J., and A. Brisse. 2010. High-Temperature    Electrolysis Using Planar Solid Oxide Fuel Cell Technology: A    Review. IN: Hydrogen Energy. Edited by Detlef Stolten. Wiley-VC    Verlag GmbH & Co. KGaA, Weinheim. ISBN: 978-3-527-32711-9. pp.    227-242.-   19. ITEA. 2010. ISOTHERM Pwr® Flameless Oxy-Combustion technology    (see website: http://www.iteaspa.it/technologies.asp). Website    accessed July 2010.-   20. Norman, J. H. , Besenbruch, G. E. , O'Keefe, D. R., and C. L.    Allen. Undated. “Thermochemical Water-Splitting Cycle, Bench-Scale    Investigations, and Process Engineering, Final Report for the Period    February 1977 through Dec. 31, 1981,”General Atomics Report    GA-A16713, DOE Report DOE/ET/26225-1. IN: Brown, L. C., Lentsch, R.    D., Besenbruch, G. E., Schultz, K. R., and J. E. Funk. 2003.    Alternative Flowsheets for the Sulfur-Iodine Thermochemical Hydrogen    Cycle. GA-A24266. General Atomics. San Diego, Calif. 19 pp.-   21. Kato Y, Yamashita N, Kobayashi K, Yoshizawa Y. 1996. Kinetic    study of the hydration of magnesium oxide for a chemical heat pump.    Applied Thermal Engineering Vol. 16(11). Pp. 853-62.-   22. Lewis, M. A., Masin, J. G., Vilim, R. B. 2005. “Development of    the Low Temperature Cu—Cl Thermochemical Cycle”, Proceedings of    International Congress on Advances in Nuclear Power Plants (ICAPP    '05), 15-19 May 2005, Seoul, Korea, paper 5425. IN: Chukwu, C. C,    Naterer, G. F. and M. A. Rosen. 2008 PROCESS SIMULATION OF    NUCLEAR-BASED THERMOCHEMICAL HYDROGEN PRODUCTION WITH A    COPPER-CHLORINE CYCLE. University of Ontario Institute of    Technology. 9 pp.-   23. Lewis M A. 2007. Cu—Cl cycle R&D recent research results for the    hydrolysis reaction sensitivity studies. Cu—Cl cycle research and    development at the Argonne National Laboratory. Canadian hydrogen    workshop on hydrogen production from non-fossil Sources. Oshawa,    Ontario, Canada: University of Ontario Institute of Technology; Dec.    20, 2007.-   24. Li, J., Suppiah, S. 2007. “Recent Advances in Nuclear Hydrogen    Research activities at AECL”, Presentation at ORF Workshop, 28 May    2007, University of Ontario Institute of Technology, Oshawa,    Ontario, Canada. IN: Chukwu, C. C, Naterer, G. F. and M. A. Rosen.    2008 PROCESS SIMULATION OF NUCLEAR-BASED THERMOCHEMICAL HYDROGEN    PRODUCTION WITH A COPPER-CHLORINE CYCLE. University of Ontario    Institute of Technology. 9 pp.-   25. Mogensen, M., Jensen, S. H., Hauch, A., Chorkendorff, I. And T.    Jacobsen. 2006. Proceedings of the 7^(th) Lucerne Fuel Cell Forum    (ed., U. Bossel). 3-7 Jul. 2006. Lucerne, P0301. IN: Mohsine, Z.,    Schefold, J., and A. Brisse. 2010. High-Temperature Electrolysis    Using Planar Solid Oxide Fuel Cell Technology: A Review. IN:    Hydrogen Energy. Edited by Detlef Stolten. Wiley-VC Verlag GmbH &    Co. KGaA, Weinheim. ISBN: 978-3-527-32711-9. pp. 227-242.-   26. Mohsine, Z., Schefold, J., and A. Brisse. 2010. High-Temperature    Electrolysis Using Planar Solid Oxide Fuel Cell Technology: A    Review. IN: Hydrogen Energy. Edited by Detlef Stolten. Wiley-VC    Verlag GmbH & Co. KGaA, Weinheim. ISBN: 978-3-527-32711-9. pp.    227-242.-   27. Naterer, G. F. , 2008. Second Law viability for upgrading waste    heat for thermochemical hydrogen production. International Journal    of Hydrogen Energy. Vol. 33. Pp. 6037-6035.-   28. Noglik, A., Roeb, M., Rzepczkyk, T., Hinkley, J., Sattler, C.,    and P. Pitz-Paal. 2009. Solar Thermiochemical Genreation of    Hydrogen: Development of a Receiver reactor for the Decompostion of    Sulfuric Acid. J Solar Energy Eng. 131, 011003-1-011003-7. IN:    Sattler, C., 2010. Thermochemical Cycles. IN: Hydrogen Energy.    Edited by Detlef Stolten. Wiley-VC Verlag GmbH & Co. KGaA, Weinheim.    ISBN: 978-3-527-32711-9. pp. 189-206.-   29. Praxair. 2010. Oxycoal combustion website.    http://www.praxair.com/praxair.nsf/AllContent/A788B0554A83B3ED852572A000598    AFA?OpenDocument&URLMenuBranch=73BE13303189965D8525735B0064CB7C .    Accessed July 2010.-   30. Rubin, E. S., Rao, A. B., and M. B. Berkenpas. 2007. Technical    Documentaiton: Oxygen-based Combustion Systems (Oxhyfuels) with    Carbon Capture and Storage (CCS). Carnegie Institute of Technology.    Pittsburgh, Pa. 47 pp.-   31. Sattler, C., 2010. Thermochemical Cycles. IN: Hydrogen Energy.    Edited by Detlef Stolten. Wiley-VC Verlag GmbH & Co. KGaA, Weinheim.    ISBN: 978-3-527-32711-9. pp. 189-206.-   32. Shah, M. 2006. Oxy-fuel Combustion for CO2 Capture from PC    Boilers. Praxair, Inc. Tonawanda, N.Y. 8 pp.-   33. Spoelstra S, Haije W G, Dijkstra J W. 2002. Techno-economic    feasibility of high-temperature high-lift chemical heat pumps for    upgrading industrial waste heat. Applied Thermal Engineering.    Vol. 22. Pp. 1619-30.-   34. U.S. Department of Energy. 2004. Waste Heat Reduction and    Recovery for Improving Furnace Efficiency, Productivity and    Emissions Performance. DOE/GO-102004-1975. Industrial Technologies    Program. Washington, D.C. 10 pp.-   35. U.S. Department of Energy. 2006. Energy Tips—Process Heating,    Process Heating Tip Sheet #10. January 2006. USDOE Energy Efficiency    and Renewable Energy, Industrial Technologies Program. Washington,    D.C. 2 pp.-   36. Ushikoshi, K., Mori, K., Kubota, T., Watanabe, T., and M.    Saito. 2000. Methanol Synthesis from CO2 and H2 in a Bench-Scale    Test Plantg. Appli. Organometal. Chem. 14. Pp. 819-825.-   37. Wall T., Gupta R., Buhre B., Khare S. Oxy-fuel (O2/CO2, 02/RFG)    technology forsequestration-ready CO2 and emission compliance. The    30th international technical conference on coal utilization & fuel    systems, coal technology: yesterday-todaytomorrow, Clearwater, Fla.,    USA, 2005. IN: Hong et al. (2008)-   38. Wang, Z. L. , Naterer, G. F. , Gabriel, K. S. , Gravelsins, R.    and V. N. Daggupati. 2009. Comparison of different copper-chlorine    thermochemical cycles for hydrogen production. International Journal    of Hydrogen Energy (34 (2009). Pp 3267-3276.-   39. Wang, Z. L. , Naterer, G. F. , Gabriel, K. S. , Gravelsins, R.,    and V. N. Daggupati. 2010. Comparison of sulfur-iodine and    copper-chlorine thermochemical hydrogen production cycles.    International Journal of Hydrogen Energy. 35 (2010). Pp. 4820-4830.-   40. Werner, R. H., ed. “Synfuels from Fusion—Using the Tandem Minor    Reactor and a Thermochemical Cycle to Produce Hydrogen,” Lawrence    Livermore Laboratory Report UCID-19609, Nov. 1, 1982. IN: Brown, L.    C., Lentsch, R. D., Besenbruch, G. E., Schultz, K. R., and J. E.    Funk. 2003. Alternative Flowsheets for the Sulfur-Iodine    Thermochemical Hydrogen Cycle. GA-A24266. General Atomics. San    Diego, Calif. 19 pp.-   41. White, V. and K. Fogash. 2009. Purification of Oxy-fuel-Derived    CO2: Current Developments and Future Plans. Presented at 1^(st)    Oxy-fuel Combustion Conference. Sep. 8-11, 2009, Cottbus, Germany.-   42. Wu, S., Kukoski, A., Jin, P., Tigges, K. D., Klauke, F., Bergins    C., Kuhr C., and S. Rehfeldt. 2009. Development of Oxy-fuel    Combustion Technology for Existing Power Plants. Unpublished    literature. Hitachi Power Systems America, Ltd. Basking Ridge, N.J.    Accessed from www.hitachipowersystems.us, December 2009. 10 pp.-   43. Zhou, W. and D. Moyeda. 2010. Process Evaluation of Oxy-Fuel    Combustion with Flue Gas Recycle in a Conventional Utility Boiler.    Energy Fuels. Vol. 24. Pp. 2162-2169.

EXEMPLIFICATION

An example of how the recuperative combustion system integrates with theISOTHERM® pressurized coal combustion follows. The ISOTHERM® system, asdesigned, makes use of an ASU, which requires nearly 20% of grossfacility power output to operate. The recuperative combustion systemdescribed herein is integrated with this system, replacing the ASU witha water splitter, and resulting in significant reductions in powerpenalty for the plant. FIG. 2 provides a system schematic of theintegration of the recuperative combustion system with the ISOTHERM®oxy-fuel (coal) combustion process. Table 1 provides the temperature andmass data from each process, and includes a description of each stepnumerically keyed to the process step numbers in FIG. 2. Systemefficiencies, heat requirements and electrical requirements may varywith the choice of water splitter employed (i.e., Cu—Cl Cycle, S—ICycle, HyS Cycle, high-temperature electrolysis or other suitable watersplitter known to those in the art). In all cases ASU (and correspondingpower requirements) are significantly reduced, or eliminated, due to theproduction of oxygen from the water splitter, hence increasing systemefficiency. Additional heat and mass balance tests are required tofurther quantify power penalty reductions and corresponding increases insystem efficiencies with the recuperative combustion system relative tothe ISOTHERM® base case (i.e., where oxygen is supplied for combustionfrom an ASU).

TABLE 1 Recuperative Combustion System Integration with ISOTHERM PWRSystem for a 875 MW_(TH) Pressurized Coal Fired Power Plant Mass FlowPressure Temperature Rate State # Process Stream State/Description (bar)(C.°) (kg/s) 2-1 Condensate leaves the condenser and is compressed 0.4132.6 198.1 by the first feedwater pump. 2-2 The pressurized condensateleaves the first feedwater 11.2 32.9 198.1 pump and enters the AcidCondenser Unit where most of the latent enthalpy in the flue gases isrecovered. 2-3 Condensate leaves the Acid Condenser Unit and picks 11.2158.7 198.1 up more thermal energy from the Coal Combustor Unit walls.2-4 Condensate leaves the Coal Combustor Unit walls and 11.2 177.2 198.1enters the Deaerator Unit. 2-5 According to the saturation condition ofthe deaerator, 11.2 179.9 210 the design point pressure level fixes theexit temperature of the water leaving the deaerator. After thedeaerator, the feedwater stream at state 2-5 is pumped to thesupercritical state, by the second feedwater pump. 2-6 After leaving thesecond feedwater pump, the 250 215 210 supercritical steam feedwater isheated regeneratively and enters the Heat Recovery Steam Generator(HRSG). 2-7, The supercritical steam feedwater leaves the Heat 250 600210 2-9 to 2-12 Recovery Steam Generator (HRSG) where it was (processheated to 600° C. at 250 bars. It then enters the High stream/ PressureTurbine (HPT) to generate electricity. state Resultant subcritical steamleaves the HPT and is description circulated for 1) steam injection intothe Coal only) Combustor Unit (State 2-8); 2) reheat toHRSG/Intermediate-Pressure Turbine (IPT; States 2-9 and 2-10), andreheat to HRSG/Low-Pressure Turbine (LPT; States 2-11 and 2-12); and 3)regeneratively reheated to supercritical steam as in State 2-6 for entryto HRSG followed by the HPT. 2-8 Steam is bled from the high pressuresteam turbine to 70 316 3 be injected into the pressurized combustor inorder to atomize the slurry particles. 2-13 An oxygen stream is producedfrom the Water 10 201.2 73.52 Splitter Section. Produced oxygen is mixedwith the recycled flue gases, state 2-19. 2-14 The recycled fluegas/oxygen mixture is injected 10 256.5 335.8 into the pressurized CoalCombustor. 2-15 The Coal Combustor Unit yields flue gases at about 101549.7 382.8 1550° C. Some of the thermal energy in the flue gas istransferred through one or more heat exchangers to the Water SpltterSection. The flue gas temperature downstream of the heat exchangers isat least 800° C. 2-16 The flue gas stream leaves the Water Splitter 10800 1004.7 Section and enters the HRSG and transfers thermal energy tothe steam while being cooled down to state 2-17. 2-17 The flue gasstream leaves the HRSG and is either 1) 9.351 259.7 1004.7 recycled, or2) passed through the Acid Condenser Unit. 2-18 The flue gas stream iscooled down to 800° C. - as 10 268.7 621.9 necessary - by recycled fluegases. 2-19 Recycled gases leaving the HRSG are either used 10 268.7262.2 for 1) mixing with the oxygen stream for injection into the CoalCombustor Unit, or 2) cooling flue gas stream downstream from the WaterSplitter Section to 800° C. 2-20 A portion of the flue gas exhauststream goes to acid 9.351 259.7 120.6 condenser for cooling/heatrecovery 2-21 Cooled flue gas leaves acid condenser and enters 9.35160.51 87.7 Carbon Dioxide Purification/Compression Unit. 2-22 CO₂ ispumped to the Methanol Reactor Section. 110 30 72.5 2-23 Exhaust gasesare monitored and vented. 1.2 30 16.9 2-24 Hydrogen gas (H₂) is producedin the Water 50 25 9.96 Splitter Section and pumped to the MethanolReactor. 2-25 A methanol/water mixture is produced in the 75 250 82.46(Methanol Methanol Reactor and separated in the and Water);Methanol/Water Separation Unit through either 52.78 pervaporation ordistillation. (Methanol); 29.68 Water) 2-26 Water is separated frommethanol through 1.38 26.7 29.68 (water distillation in theMethanol/Water Separation Unit, from Methanol and is treated, asnecessary, and recirculated to the Reactor); Water Splitter Section).Additional water is 29.56 (water provided from State 2-29 (waterseparated from from MTG petrochemical distillates in MTG Unit) andUnit); 82.78 external water sources, as necessary. (total water neededto run Water Splitter Section to produce adequate H₂ and O₂ forCombustor System) 2-27 Methanol resulting from the Methanol/Water 1 2552.78 Separation Unit is pumped to bulk storage tanks for temporarystorage prior to transport and sale, or pumped to the Methanol toGasoline Reactor Unit. 2-28 Hydrocarbon distillate products of theMethanol to ≦55 (MTG 200-540 (MTG 23.22 Gasoline Reactor/DistillationSystem are pumped to Reactor Unit); Reactor Unit); bulk storage tanksfor temporary storage prior to 1 (storage) 25 (storage) transport andsale. 2-29 Water resulting from the Methanol to Gasoline 1 25 29.56Reactor/Distillation System is treated and recirculated to theWater-Splitting Section via State 2-26. 2-30 Coal/Water Slurry injectionto Coal Combuster. Coal 30 - Coal Only is supplied in the form of acoal-water slurry stream which contains 0.35 kg water per 1 kg of itstotal weight.

1. A method for oxy-fuel combustion, comprising: providing a systemcomprising a combustion chamber arranged and disposed to receive fuel,oxygen and recycled flue gas and combust said fuel, oxygen and recycledflue gas to produce heat and heated flue gas; capturing heat produced bythe oxy-fuel combustion; using a portion of the heat to power a watersplitter, thereby generating hydrogen gas and oxygen gas; andtransferring the oxygen gas from the water splitter to the combustionchamber for use in said oxy-fuel combustion.
 2. The method of claim 1,further comprising providing an air separator unit, or another externaloxygen supply, wherein the combustion chamber is arranged and disposedto receive oxygen from the air separator unit and/or external oxygensupply and/or the water splitter.
 3. The method of claim 1, furthercomprising one or more heat exchangers for the capture and transfer ofheat from the heated flue gas to the water splitter. 4-5. (canceled) 6.The method of claim 1, wherein the amount of oxygen required from theair separator unit and/or external oxygen supply, is reduced oreliminated in proportion to the amount of oxygen provided by the watersplitter.
 7. (canceled)
 8. The method of claim 1, wherein the oxy-fuelcomprises any hydrocarbon-based fuel.
 9. The method of claim 8, whereinthe oxy-fuel comprises a coal/water slurry.
 10. (canceled)
 11. Themethod of claim 1, wherein the water splitter produces hydrogen gas andoxygen gas by means of high-temperature electrolysis.
 12. The method ofclaim 1, wherein the water splitter produces hydrogen gas and oxygen gasby means of a thermochemical cycle.
 13. The method of claim 12, whereinthe thermochemical cycle is selected from: a hybrid copper-chlorinecycle; a sulfur-iodine cycle; and a hybrid sulfur cycle. 14-15.(canceled)
 16. The method of claim 13, wherein the thermochemical cycleis a hybrid copper-chlorine cycle.
 17. The method of claim 16, whereinthe hybrid copper-chlorine cycle is selected from: a 3-step cycle, a4-step cycle, and a 5-step cycle.
 18. The method of claim 17, whereinthe hybrid copper-chlorine cycle is the 4-step cycle.
 19. The method ofclaim 18, wherein the 4-step cycle of the hybrid copper chlorine cyclemay be represented by the following steps: Step I:Cu(s)+2HCl(g)→2CuCl(molten)+H₂(g) Step II:4CuCl(s)→2Cu(s)+2CuCl₂(aq)+HCl(aq) Step III:CuCl₂(aq)+n_(f)H₂O(l)→CuOCuCl₂(s)+2HCl(g)+(n_(f)−1)H₂O(g) Step IV:CuOCuCl₂(s)→2CuCl(molten)+0.50₂(g)
 20. (canceled)
 21. The method ofclaim 19, wherein n_(f) is 5-30.
 22. (canceled)
 23. The method of claim1, wherein the hydrogen from the water splitter is used directly orindirectly in a subsequent process.
 24. (canceled)
 25. The method ofclaim 1, wherein the hydrogen from the water splitter and the carbondioxide from the combustion flue gas are reacted to form methanol andwater. 26-29. (canceled)
 30. The method of claim 1, wherein the amountof oxygen that the combustion chamber requires from an air separatorunit and/or external oxygen supply, is reduced or eliminated inproportion to the amount of oxygen received from the water splitter,resulting in increased efficiency.
 31. The method of claim 30, whereinthe increased efficiency is measured in terms of increased gross poweroutput of the combustion process.
 32. The method of claim 31, whereinthe gross power output of the combustion process is increased by 1-20%.33-35. (canceled)
 36. An oxy-fuel combustion system, comprising: acombustion chamber arranged and disposed to receive fuel, oxygen andrecycled flue gas and combust said fuel, oxygen and recycled flue gas toproduce heat and heated flue gas containing carbon dioxide; one or moreheat exchangers arranged and disposed to capture heat produced by theoxy-fuel combustion and transfer said heat to a water splitter; and awater splitter, for the conversion of heat and electricity into hydrogengas and oxygen gas.
 37. The system of claim 36, further comprising anair separator unit, or another external oxygen supply, wherein thecombustion chamber is arranged and disposed to receive oxygen from theair separator unit, or external oxygen supply, and/or the watersplitter.
 38. The system of claim 36, wherein one or more heatexchangers serve as the means to capture the heat produced by theoxy-fuel combustion and to transfer said heat to the water splitter.39-62. (canceled)